Process for producing aromatic amines by gaseous phase hydrogenation

ABSTRACT

A process for the production of aromatic amines by catalylic hydrogenation.

FIELD OF THE INVENTION

Process for the production of aromatic amines by gas phase hydrogenationThe present invention relates to a process for the production ofaromatic amines by catalytic hydrogenation of corresponding aromaticnitro compounds in the gas phase on inunobilised supported catalysts.

BACKGROUND OF THE INVENTION

It is known to hydrogenate nitrobenzene and other nitroaromatics toyield the corresponding aromatic amines in the gas phase on immobilisedsupported catalysts. FR 25 25 591, for example, thus describes a processfor hydrogenating nitrobenzene on stationary copper catalysts. DE-A 2135 155 and 2 244 401 futhermore describe processes for reducing nitrocompounds in the presence of supported catalysts containing palladiumusing spinels as the support materials. DE-A 2 849 002 also describes aprocess for the catalytic hydrogenation of nitrobenzene, in whichhydrogenation is performed in the presence of a multi-componentsupported catalyst. A disadvantageous feature of the gas phasehydrogenations described in the stated patent publications is the lowloading (specific loading) of the catalysts. The stated or calculatedloadings vary between 0.2 and 0.6 kg/l×h. Loading is here defined as thequantity of nitroaromatics in kg which are reacted per litre of catalystbed within one hour. The low catalyst loading is accompanied by anunsatisfactory space/time yield in large scale industrial processes forthe production of aromatic amines.

DE-A 4 039 026 describes a gas phase hydrogenation on palladiumcatalysts which may be highly loaded. Catalyst loading in this processis between 0.6 and 0.95 kg/l×h. Industrial performance of the processhas, however, revealed that after only a short time conversion of thenitrobenzene proceeds only incompletely, such that the condensatecontains not inconsiderable quantities of unreacted nitrobenzene inaddition to the aromatic amine which has been formed. If thenitrobenzene content is to be reduced to below the tolerable limit, theaniline must be purified by elaborate purification methods(distillation). It has furthermore been found that, at the high catalystloading, the catalysts used in DE-A 4 039 026 have only anunsatisfactory service life.

The object of the present invention was to provide a process for theproduction of aromatic amines by catalytic hydrogenation of thecorresponding nitro compounds in the gas phase, which may be performedwithout problems on an industrial scale and which ensures an elevatedspace/time yield, combined with improved productivity of the catalystsused.

This objective is achieved with the process according to the invention.

SUMMARY OF THE INVENTION

The present invention accordingly provides a process for the productionof aromatic amines by catalytic hydrogenation of corresponding aromaticnitro compounds in the gas phase on immobilised catalysts which containone or more metals of groups VIIIa, Ib, IIb, IVa, Va, VIa, IVb and Vb ofthe periodic system of elements (Mendeleyev) on a ceramic supportmaterial having a BET surface area of less than 40 m² /g, at molarratios of hydrogen to nitro group or nitro groups of 3:1 to 30:1 andtemperatures in the catalyst bed of 180 to 500° C., which process ischaracterised in that loading of the catalyst with the aromatic nitrocompounds used is increased continuously or step-wise from 0.01-0.5 to0.7-5.0 kg/l×h, wherein maximum loading is achieved within 10 to 1000hours.

DESCRIPTION OF THE INVENTION

Aromatic nitro compounds which may be used in the process according tothe invention are those of the following formula ##STR1## in which R¹and R² are identical or different and denote C₁ -C₄ alkyl, in particularmethyl and ethyl, hydrogen or the nitro group.

Nitrobenzene or the isomeric nitrotoluenes, in particular nitrobenzene,are used in the hydrogenation process according to the invention.

Catalysts suitable for the process according to the invention are anyknown for the hydrogenation of nitro groups and in which the metals ofthe above-mentioned main and sub-groups of the periodic system ofelements have been applied onto a ceramic support. The metals may havebeen applied onto the catalyst support in elemental form or in the formof a compound. Metals which may be applied onto a ceramic support whichmay in particular be mentioned are Fe, Co, Pd, Pt, Cu, Ag, Zn, Ti, Zr,Hf, V, Nb, Ta, Cr, Mo, W, Ge, Sn, Pb, As, Sb, Bi, particularlypreferably Fe, Pd, Pt, Cu, Ti, V, Nb, Cr, Mo, Sn, Pb, Sb, Bi, veryparticularly preferably Pd, V, Pb, Bi. The metals may have been appliedonto the ceramic support individually or as a mixture with each other.

Suitable ceramic support materials for the stated metals are inprinciple any ceramic solids having BET surface areas of less than 40 m²/g, preferably of less than 20 m² /g, in particular of less than 10 m²/g. Suitable ceramic solids are in particular metal oxides and/or metalmixed oxides of the elements magnesium, aluminium, silicon. germanium,zirconium and titanium, with a-aluminium oxide being preferably used asthe support material.

Catalysts as are described in DE-A 2 849 002 have proved to beparticularly suitable. Catalysts which may in particular be mentionedare those containing primarily palladium, vanadium and lead among othermetals deposited as a shell on a-aluminium oxide. Particular emphasisshould thus be placed on catalysts which contain a) 1 to 50 g ofpalladium, b) 1 to 50 g of titanium, vanadium, niobium, tantalum,chromiun, molybdenum and/or tungsten and c) 1 to 20 g of lead and/orbismuth per litre of oxide support material, wherein the supportmaterial has a surface area of less than 40, preferably of less than 20,in particular of less than 10 m² /g. Of the stated metals in the mixedcatalyst, palladium, vanadium and lead in the above-stated quantitieshave again proved particularly favourable in use.

Production of the catalysts is known from the cited patent literature.It has been found to be advantageous if the active component of thecatalyst is deposited as a well defined zone as close as possible to thesurface of the moulded catalyst shape and the interior of the supportmaterial contains no metal. The catalysts may here be produced with orwithout pretreatment of the support material with a base.

In principle, the supported catalysts for the process according to theinvention may have any desired shape, such as spheres, rods, Raschigrings, pellets or tablets. The moulded shapes used are preferably thosewhich, as beds, exhibit low flow resistance with good gas surfacecontact, such as Raschig rings, saddles, waggon wheels and/or spirals.It is possible in the process according to the invention for thecatalyst bed in the reactors to consist solely of the supported catalystor for it to be additionally diluted with inert support material orother inert packing, such as for example glass or ceramic. The catalystbed may consist of up to 90 wt. %, preferably up to 75 wt. %, inparticular up to 50 wt. % of inert packing or support material. The bedmay have a dilution gradient such that dilution decreases in thedirection of flow. At the feed surface, the bed may contain between 10to 50 wt. % of packing material and, at the outlet end, consist of 80 to100 wt. % of pure supported catalyst.

Reactors which may be used for the process according to the inventionare any known reactors which are suitable for gas phase reactions withcooled, stationary catalyst beds. Suitable reactors are, for example,multi-tube reactors in which the catalyst is located within tubes aroundwhich a heat transfer medium flows and reactors in which, conversely,the heat transfer medium flows within the tubes and the catalyst islocated outside the tubes. Such reactors are known, for example, fromDE-A 2 848 014 and 3 007 202.

Reactors which have proved to be particularly advantageous for theprocess according to the invention are those in which the heat transfermedium flours within the tubes and the catalyst is located outside thetubes (Linde type reactors). In comparison with conventional multi-tubereactors, constant operating times between regenerations are observedover many regeneration cycles in these reactors.

The length of the catalyst bed in the direction of flow in the processaccording to the invention is from 0.5 to 20, preferably from 1 to 10,more preferably from 2 to 6 m.

The process according to the invention is preferably performed at molarratios of hydrogen to nitro group or nitro groups of the nitro compoundused of 4:1 to 20:1, preferably from 5:1 to 10:1.

Hydrogen concentration may be reduced by incorporating inert carriergases, such as nitrogen, helium, argon and/or steam. Nitrogen ispreferably used. Up to 10 mol, preferably up to 3 mol, more preferablyup to 1 mol of inert carrier gas may be introduced per mol of hydrogen.

Dilution with inert carrier gas is preferably used at the beginning ofthe period of operation with fresh catalyst and after regeneration ofthe catalyst by burning off with air and reduction with hydrogen.Dilution with inert gas is preferably performed for the first 300 hours,more preferably for the first 200 hours, most preferably for the first100 hours after starting up again with inert gas.

Deactivated beds are regenerated with N₂ /O₂ mixtures at temperatures ofbetween 200 and 400° C., preferably between 250 and 350° C. on the solidcatalyst, with regeneration preferably being begun at N₂ contents ofbetween 90 and 99% in the gas stream and the O₂ content being raised instages to pure air during burning off, and, at the end of regeneration,tenacious carbon residues may optionally be burnt off with pure oxygen.Inert carrier gases other than nitrogen, such as for example argon,helium or steam, may also be added to oxygen or air.

The process according to the invention is preferably performed attemperatures within the range from 180 to 500, more preferably from 200to 460, most preferably 220 to 440° C. It may be advantageous in thisconnection if the temperature of the cooling medium in the processaccording to the invention is raised continuously or step-wise duringthe operating cycle.

The process according to the invention is operated at a pressure of 0.5to 5, preferably of 1 to 3 bar.

An essential feature of the process according to the invention is thatloading of the catalyst with the aromatic nitro compounds used is raisedcontinuously or step-wise to the desired loading value within a certainperiod of time. A mode of operation is accordingly preferred in whichcatalyst loading is raised continuously or step-wise within 20 to 500,more preferably within 30 to 300, most preferably within 40 to 200hours, and from 0.01-0.5, preferably from 0.1-0.4, more preferably from0.15-0.3 to 0.7-5.0, preferably to 0.8-3.0, more preferably to 1.0-2.0kg/l×h.

The elevated final loading is maintained until unreacted educt breaksthrough. Once the educt concentration becomes excessively high at theend of the reactor, the temperature of the heat transfer medium may beraised and/or educt loading reduced, in order to delay an interruptionin production for catalyst regeneration.

In a particular embodiment of the process according to the invention,beds are used in which the supported catalysts described above are usedmixed with another solid catalyst, in which palladium has been appliedonto graphite or coke containing graphite as the support material and inwhich the support material has a BET surface area of 0.05 to 10 m² /g,preferably of 0.2 to 5 m² /g. The palladium content of thegraphite-supported catalyst is 0.1 to 7 wt. %, preferably 1.0 to 6 wt.%, more preferably 1.5 to 5 wt. %, most preferably 2 to 4 wt. %. Thestated supported palladium catalysts containing graphite are described,for example, in DE-A 2 849 002.

If the supported catalysts based on a ceramic support material and thesupported catalysts based on graphite as the support material are usedas mixtures with each other, the ratio of weights in the mixture is 1/lto 1000/l, preferably 10/l to 100/l, more preferably 90/1 to 99/1(ceramic supported catalyst to graphite supported catalyst).

The graphite catalyst is preferably used in the first third of thecatalyst bed, but may also be distributed uniformly throughout theentire bed.

The process according to the invention may, for example, be realisedindustrially in the following manner: a circulating gas streamsubstantially consisting of hydrogen and a little water is compressed inorder to overcome the plant's resistance to flow. The gas stream isheated by countercurrent heat exchange, wherein the heat is taken, forexample, from the circulating gas stream before condensation of theproducts. The circulating gas stream is adjusted to the desiredtemperature. The nitroaromatic compound to be hydrogenated is vaporisedand superheated in fresh hydrogen, which replaces that consumed, and thetwo gas streams are then mixed. The gas mixture is introduced into atemperature-controlled reactor with stationarily arranged catalyst. Theliberated heat of reaction is removed from the reactor by means of aheat transfer medium. The product stream leaving the reactor is used toheat the circulating gas stream and then cooled with water until theaniline formed condenses. The liquids are discharged, together with asmall quantity of circulating gas in order to remove gases, for examplenitrogen, which would otherwise accumulate. The circulating gas is thenreturned to the compressor.

In a preferred embodiment of the process according to the invention, thecatalyst bed is introduced into a Linde type reactor (heat transfermedium flows within the reactor tubes and the catalyst is arrangedoutside the heat transfer medium tubes) and the process is performed asdescribed above. To this end, fresh or freshly regenerated catalyst isoperated with nitrogen/hydrogen mixtures for the first hours. Theadvantages of this preferred mode of operation are: elevated selectivityand long periods of time between regenerations, even after manyproduction cycles.

The process according to the invention is in particular distinguished bythe elevated space/time yields which may be achieved, combined with areduction in size of the necessary plant and equipment and distinctlyincreased catalyst productivity. Considerable increases in output maythus be achieved in existing plant by means of the process according tothe invention.

A particularly favourable development of the process according to theinvention is that in which supported catalysts having ceramic supportmaterials are used in combination with supported catalysts having agraphite support, using a reactor type in which the supported catalystsare located outside the tubes containing the heat transfer oil (Lindetype).

The invention is further illustrated in the following examples. Allreferences to parts and percentages are by weight unless otherwiseindicated.

EXAMPLES Catalyst Preparation on a Ceramic Support. Example 1

One litre of an a-Al₂ O₃ support in spherical form having a diameter of3 to 5 nmm, a BET surface area of 9.8 m² /g, an absorption capacity of45.1 ml of water per 100 g of support and a bulk density of 812 g/l wasimpregnated with 366 ml of an aqueous solution containing 10.8 g(corresponding to 0.27 gram-equivalents) of NaOH. The solution wascompletely absorbed by the support within a few minutes.

The moist support was dried in a warm, increasingly strong stream ofair. Drying time to constant weight was approx. 15 minutes. Aftercooling, the residual moisture content was approx. 1% of the absorptioncapacity of the support.

The dry support pretreated in this manner was saturated in accordancewith its absorption capacity with 366 ml of an aqueous sodiumtetrachloropalladate solution which contained 9 g of palladium(corresponding to 0.169 gram-equivalents) and left to stand for 15minutes. In order to reduce the palladium compound which had beendeposited on the support to metallic palladium, the catalyst was coveredwith 400 ml of a 10% aqueous hydrazine hydrate solution and left tostand for 2 hours. The catalyst was then thoroughly rinsed withcompletely deionised water until no ions of the compounds used incatalyst production were any longer detectable in the rinsing water,which took some 10 hours.

Drying to constant weight was then again performed in a strong,increasingly warm stream of air.

The catalyst containing Pd was then saturated with 366 ml of an aqueoussolution containing 9 g of vanadium as vanadyl oxalate. The support wasdried in a stream of warm air as described above. The catalyst was thenheat treated for 6 hours at 300° C. in a tubular oven, wherein theoxalate was decomposed.

The catalyst was then saturated with 366 ml of an aqueous solutioncontaining 3 g of lead as lead acetate and again dried in the increasingair stream.

The completed catalyst contained 9 g of Pd, 9 g of vanadium and 3 g oflead and matched the catalyst from Example 1 in DE-OS 2849002.

Examples with Catalyst Beds Containing Ceramic Supported CatalystsExample 2 (Comparative Example 1)

A 285 cm deep bed of a catalyst produced according to Example 1 wascharged into a reaction tube of an internal diameter of approx. 26 mmand temperature-controlled with oil. The catalyst was flushed first withnitrogen, then with hydrogen and then heated within 5 hours to 240° C.in a stream of hydrogen of approx. 770 Nl/h. Vaporisation ofnitrobenzene in the stream of hydrogen was then begun. Thenitrobenzene/hydrogen mixture arrived at the face of the catalyst bed atapprox. 230° C. The specific loading of the catalyst was increasedwithin 48 hours from 0.1 to 0.47 kg/l×h, corresponding to a loading perunit area of 1425 kg/m² ×h, such that a mean loading of 0.46 kg/l×h wasachieved. The temperature in the entire bed was monitored throughout thetest and it was ensured that the catalyst temperature did not exceed440° C. at any point.

The oil temperature was raised in 20° C. stages after approx. 700, 800and 900 hours from 240° C. to 300° C. The variation in oil temperaturealong the reaction tube was approx. ±1° C. The flow velocity of the oilalong the tube surface was approx. 1.5 m/s.

The catalyst achieved a service life of approx. 1450 hours, after whichthe nitrobenzene content of the condensate rose from 0 to approx. 300ppm. Thereupon, the catalyst had to be regenerated by burning off. Underthe stated reaction conditions, the catalyst thus achieved aproductivity of approx. 670 kg/l. Mean selectivity was 98.4%.

The catalyst behaved identically in the second cycle after regeneration,giving a service life of 1400 hours and 98.6% selectivity.

Example 3

A 285 cm deep bed of a catalyst produced according to Example 1 wascharged into a reaction tube of an internal diameter of approx. 26 mmand temperature-controlled with oil. The catalyst was flushed first withnitrogen, then with hydrogen and then heated within 5 hours to 240° C.in a stream of hydrogen of approx. 1528 Nl/h. Vaporisation ofnitrobenzene in the stream of hydrogen was then begun. Thenitrobenzene/hydrogen mixture arrived at the face of the catalyst bed atapprox. 230° C. The specific loading of the catalyst was increasedwithin 80 hours from 0.2 to 1.05 kg/l×h, corresponding to a loading perunit area of 2994 kg/m² ×h, such that a mean loading of 1.03 kg/l×h wasachieved. It was ensured that the catalyst did not become hotter than440° C. at any point throughout the test.

The oil temperature was raised in 20° C. stages after approx. 700, 800and 900 hours from 240° C. to 300° C. The variation in oil temperaturealong the reaction tube was approx. ±1° C. The flow velocity of the oilalong the tube surface was approx. 1.5 m/s.

The catalyst achieved a service life of approx. 1050 hours, after whichthe nitrobenzene content of the condensate rose from 0 to approx. 300ppm, whereupon the catalyst had to be regenerated by burning off. Underthe stated reaction conditions, the catalyst thus achieved aproductivity of approx. 1080 kg/l, approx. 1.6 times as productive as inthe Comparative Example.

Mean selectivity was 99.0%, 0.6% higher than that of the Comparativetest.

The catalyst behaved identically in the second cycle after regeneration,giving a service life of 990 hours and 99.2% selectivity.

Since the catalyst regeneration time of a few hours is not of greatsignificance in comparison with the cycle time, the quantity of anilineproduced per unit time is twice as high as in the Comparative Exampledue to the elevated loading.

Catalyst preparation on a graphite support.

The support material used comprised EG 17 graphite pellets supplied byRingsdorff having a BET surface area of approx. 0.4-0.8 m² /g. Grainsize was between 4 and 5 mm.

The Examples should not in any way be viewed as restrictive: similarresults are also achieved with other graphites and materials containinggraphite having a low BET surface area.

The catalysts are prepared in the following manner: EG 17 graphitepellets supplied by Ringsdorff (4-5 mm pellets, shaken density 650-1000g/ml) having an absorption capacity of 7 ml of acetonitrile per 10 g ofsupport are introduced into a rotatable vessel and combined, while thevessel is rotated, with a solution of palladium acetate in acetonitrile.The mixture is kept in motion until the solution has been completelyabsorbed. The solid is then dried for 5 minutes in a increasingly strongstream of warm air at 40° C. The saturation and drying stages arerepeated until the desired quantity of palladium has been deposited.

The dried catalyst is then activated in a hot stream of hydrogen atstandard pressure.

Example 4

0.6% Pd on EG17

2000 g of support

3 saturation stages, each comprising

8.3 g of PdAc₂ in 140 g of acetonitrile, activated for 20 h at 370° C.in a stream of H₂.

Example 5

2.4% Pd on EG17

2000 g of support

10 saturation stages, each comprising

10 g of PdA_(c) in 140 g of acetonitrile, activated for 20 h at 370° C.in a stream of H₂.

Example With Catalyst Bed Containing Graphite Supported Catalyst andInert Supports. Example 6 (Comparative Example 2)

A 285 cm deep bed of a catalyst produced according to Example 4 wascharged into a reaction tube of an internal diameter of approx. 26 mmand temperature-controlled with oil. The catalyst bed was sub-dividedinto six sections approx. 47.5 cm in length, in which the catalyst waspacked homogeneously diluted with untreated support material. Dilutiondecreased in the direction of flow as follows: 97%, 94%, 88%, 60%, 60%,0%. The final section thus contained undiluted catalyst.

The bed was flushed first with nitrogen, then with hydrogen and thenheated within 5 hours to 240° C. in a stream of hydrogen of approx. 1528Nl/h. Vaporisation of nitrobenzene in the stream of hydrogen was thenbegun. The nitrobenzene/hydrogen mixture arrived at the face of thecatalyst bed at approx. 230° C. The specific loading of the catalyst wasincreased within 60 hours from 0.4 to 0.84 kg/l×h, corresponding to aloading per unit area of 2395 kg/m² ×h. It was ensured that the catalystdid not become hotter than 440° C. at any point throughout the test.

The variation in oil temperature along the reaction tube was approx. ±1°C. The flow velocity of the oil along the tube surface was approx. 1.5m/s.

The catalyst achieved a service life of approx. 70 hours, after whichthe nitrobenzene content of the condensate rose from 0 to approx. 300ppm. Under the stated reaction conditions, the catalyst had thusachieved a productivity of approx. 50 kg/l. Mean selectivity was 99.5%.

Raising the temperature of the heat transfer medium from 240 to 260 and300° C. did not counteract break-through of the nitrobenzene, whichrapidly rose to more than 10%.

Examples with Catalyst Beds Containing both Ceramic and GraphiteSupported Catalysts. Example 7

A 285 cm deep bed of catalysts produced according to Examples 4 and 1was charged into a reaction tube of an internal diameter of approx. 26nmu and temperature-controlled with oil. The bed consisted of twosections of approx. 50 and 235 cm in length.

The first zone 50 cm in length contained catalyst from Example 1 98%diluted with untreated support material. The second zone 235 cm inlength contained undiluted catalyst from Example 1.

The bed was flushed first with nitrogen, then with hydrogen and thenheated within 5 hours to 240° C. in a stream of hydrogen of approx. 1528Nl/h. Vaporisation of nitrobenzene in the stream of hydrogen was thenbegun. The nitrobenzene/hydrogen mixture arrived at the face of thecatalyst bed at approx. 230° C. The specific loading of the catalyst wasincreased within 70 hours from 0.6 to 0.9 and, after a further 140hours, to 1.06 kg/l×h, corresponding to a loading per unit area of 3081kg/m² ×h, such that a mean loading of 1.04 kg/l×h was achieved. It wasensured that the catalyst did not become hotter than 440° C. at anypoint throughout the test.

After 140 hours, the oil temperature was raised from 240° C. to 300° C.The variation in oil temperature along the reaction tube was approx. ±1°C. The flow velocity of the oil along the tube surface was approx. 1.5m/s.

The catalyst bed achieved a service life of approx. 1450 hours, afterwhich the nitrobenzene content of the condensate rose from 0 to approx.300 ppm, whereupon the catalyst had to be regenerated by burning off.Under the stated reaction conditions, the bed had thus achieved aproductivity of approx. 1508 kg/l and is thus approx. 2.3 times moreproductive than the bed in Comparative Example 1 and approx. 30 timesmore productive than the bed in Comparative Example 2.

Mean selectivity was 99.3%.

Example 8

A 285 cm deep bed of catalysts produced according to Examples 5 and 1was charged into a reaction tube of an internal diameter of approx. 26mm and temperature-controlled with oil. The bed consisted of an intimatemixture of the catalysts, comprising 3% catalyst from Example 5 and 97%catalyst from Example 1.

The bed was flushed first with nitrogen, then with hydrogen and thenheated within 5 hours to 240° C. in a stream of hydrogen of approx. 1528Nl/h. Vaporisation of nitrobenzene in the stream of hydrogen was thenbegun. The nitrobenzene/hydrogen mixture arrived at the face of thecatalyst bed at approx. 230° C. The specific loading of the catalyst wasincreased within 280 hours from 0.25 to 1.07, corresponding to a loadingper unit area of 3110 kg/m² ×h, such that a mean loading of 0.98 kg/l×hwas achieved. It was ensured that the catalyst did not become hotterthan 440° C. at any point throughout the test.

The oil temperature was raised in 20° C. stages after approx. 460, 1000and 1100 hours from 240° C. to 300° C. The variation in oil temperaturealong the reaction tube was approx. ±1° C. The flow velocity of the oilalong the tube surface was approx. 1.5 m/s.

The catalyst achieved a service life of approx. 1400 hours, after whichthe nitrobenzene content of the condensate rose from 0 to approx. 300ppm, whereupon the catalyst had to be regenerated by burning off. Underthe stated reaction conditions, the bed had thus achieved a productivityof approx. 1372 kg/l and is approx. 2 times more productive than the bedin Comparative Example 1 and approx. 27 times more productive than thebed in Comparative Example 2.

Mean selectivity was 99.0%.

Comparison between multi-tube reactor and Linde reactor.

Example 9 (Multi-tube reactor)

The test from Example 3 was repeated in a multi-tube reactor containing55 tubes. The tubes had an internal diameter of approx. 26 mm, a wallthickness of approx. 2 mm and a length of approx. 3.5 m.

An approx. 285 cm deep catalyst bed was charged into each tubularreactor, giving a total of approx. 83 litres. The remainder of the bedconsisted of inert packing. The total bed face area was approx. 292 cm².

Example 3 was repeated over 10 production cycles in this reactor.Regeneration was performed for 10 to 15 hours with N₂ /air mixture andwith pure air at a heat transfer medium temperature of 290 to 320° C.Mean selectivities were between 99.0 and 99.5%. The service lives forthe individual production cycles are shown in Table 1.

Example 10 (Linde reactor)

The test from Example 3 was repeated in a multi-tube reactor containing55 tubes. The tubes had an internal diameter of approx. 26 mm, a wallthickness of approx. 2 mm and a length of approx. 3.5 m. The reactor hadan internal diameter of approx. 295 mm.

The catalyst bed of approx. 83 litres was placed to a depth of approx.285 cm into the interstices between the tubes and had a face area ofapprox. 292 cm².

Wire mesh ring packing was placed above and below the catalyst bed. Thecooling medium was passed through the tubes.

Example 3 was repeated over 10 production cycles in this reactor. Meanselectivities were between 99.1 and 99.4%.

The service lives for the individual production cycles are shown inTable 1.

                  TABLE 1                                                         ______________________________________                                        Service lives of the production cycles in Examples 9 and 10 in hours.           Production cycle                                                                          Service life (Example 9)                                                                      Service life (Example 10)                       ______________________________________                                        1         1010            950                                                   2 970 930                                                                     3 940 910                                                                     4 910 900                                                                     5 870 890                                                                     6 840 880                                                                     7 810 870                                                                     8 750 860                                                                     9 710 850                                                                     10  650 840                                                                 ______________________________________                                    

Although the invention has been described in detail in the foregoing forthe purpose of illustration, it is to be understood that such detail issolely for that purpose and that variations can be made therein by thoseskilled in the art without departing from the spirit and scope of theinvention.

What is claimed is:
 1. A process for the production of aromatic aminesby catalytic hydrogenation of corresponding aromatic nitro compounds inthe gas phase on immobilised catalysts which contain a metal componentcomprising a metal selected from the group consisting of metals ofVIIIa, Ib, IIb, IVa, Va, VIa, IVb and Vb groups of the periodic systemof elements (Mendeleyev) on a ceramic support material having a BETsurface area of less than 40 m² /g, at molar ratios of hydrogen to nitrogroup or nitro groups of 3:1 to 30:1 and temperatures in the catalystbed of 180 to 500° C., wherein loading of the catalyst with the aromaticnitro compounds used is increased continuously or step-wise from0.01-0.5 to 0.7-5.0 kg/l×h, wherein maximum loading is achieved within10 to 1000 hours.
 2. A process for the production of aromatic aminesaccording to claim 1, wherein the process is performed in reactors inwhich the heat transfer medium is located within the tubes and thecatalyst is arranged outside the tubes containing the heat transfermedium.
 3. A process for the production of aromatic amines according toclaim 1, wherein the catalyst used is a supported catalyst, in whichpalladium had been applied onto graphite or coke containing graphite asthe support, which has a BET surface area of 0.05 to 10 m² /g, and thepalladium content of which catalyst is 0.1 to 7 wt. %, relative to theweight of the catalyst, in combination with a catalyst which containsone or more metals comprising a component selected from the groupconsisting of VIIIa, Ib, IIb, IVa, Va, VIa, IVb and Vb groups on aceramic support material having a BET surface area of less than 40 m²/g, and wherein the quantity ratio of graphite supported catalyst toceramic supported catalyst is between 1/1 and 1/1000 parts by weight. 4.A process for the production of aromatic amines according to claim 1,wherein the catalyst beds are diluted with inert packing and optionallyhave activity gradients.